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Microbial asparaginase recovery by membrane processes
Journal
of Membrane
Elsevier
Science
MICROBIAL
M.S.
21(1984)
B.V.,
Amsterdam
ASPARAGINASE
Hunter
Filters
Ltd.,
307 in The Netherlands
RECOVERY
Durham
BY MEMBRANE
PROCESSES*
Road, Birtley.
Co. Durham,
DH3 2SF (Great Britain)
SPARK
Microbial Technology Laboratory, Porton Down, Salisbury, Wiltshire, P.S.
307-319 - Printed
LE**
Domnick L.B.
Science,
Publishers
WARD
Domnick (Received
PHLS Centre for Applied SP4 OJG (Great Britain)
Microbiology
& Research,
and N. LADWA
Hunter
Filters Ltd.,
December
19,
1983;
Durham accepted
Road,
Birtley,
in revised
Co. Durham,
form
April
DH3 2SF (Great Britain)
12, 1984)
Summary A feasibility study in the use of membranes for the recovery of L-asparaginase from the fermentation stage has been undertaken. In the cell harvesting stage 100 1 of broth (0.55% dry weight approx) of Erwinia carotovora is concentrated to 3.2 1 by a 1 m’ prototype cross flow microfiltration unit in just over 4 hours. Cell disruption was carried out by raising the cell suspension to pH 11.5 to release the enzyme, which is stable in alkali; for a short period. After neutralisation by acetic acid addition, the enzyme is separated from the cell debris by filtration with 0.45 pm Asypor membranes. A preliminary economic analysis shows that the membrane process can compete favourably with the conventional process, which employs high speed centrifuges.
Introduction L-asparaginase is frequently used in the treatment of some forms of leukaemia [l]. This therapeutic enzyme is an intracellular product of the organism Erwinia carotouoru (NCPPB 1066). Currently the microorganism is grown in batches of 500 1 to a concentration of approximately 0.55% dry weight. Cell harvesting is carried out by a continuous flow centrifugal separator which concentrates the broth by about 36-fold. The cell paste obtained from this process is usually frozen and stored for a short period before enzyme extraction is performed. Since L-asparaginase is stable at highly alkali pH [ 11, cell disruption is conveniently carried out by the addition of alkali to the cell suspensions to pH 11.5 for about 20 minutes. The nucleic acid and other cell debris materials which are released from the disruption process form a thick, *Paper presented at IMTEC ‘83, November 8-10, 1983, **To whom all correspondence should be addressed,
0376-7388/84/$03.00
0 1984
Elsevier
Science
Sydney,
Publishers
Australia.
B.V.
308
jelly-like mass, which is removed by the addition of acetic acid, which reduces the lysate pH to 6.5. Cell debris which includes partially dissolved cell walls and aggregates of protein precipitates is removed again by centrifugation at high g forces (usually >lO,OOOg). The use of large centrifuges has many disadvantages, both from the economic and safety viewpoints. In the first instance, centrifugal installation entails very high capital costs. A good deal of supervision is required during running and the frequent bowl change is a necessary process which is tedious and labour intensive. From the safety angle, where cost cannot be measured in terms of money and where opinion can vary considerably, the operation of centrifuges can be dangerous, with the possible risk of disintegration. Specially enclosed housings and procedures must be designed to deal with pathogenic organisms. This is extremely difficult in large-scale processing. Cross flow microfiltration, particularly with the new generation of high performance asymmetric microporous membranes, should minimise the operating costs, eliminate the danger of disintegration and go a long way towards reducing the contamination hazard. Cross flow microfiltration has been used for bacterial cell harvesting for over a decade [2], mainly on a laboratory scale where only small batches of up to 10 1 of culture suspension have been concentrated to about 200 ml [3,4]. Several claims of large-scale bacterial harvesting by cross flow membrane systems have been made. In all cases, however, the final cell concentrations achieved were typically less than 2% dry weight [ 5,6]. There are several factors which limit the final concentration of the cells. Firstly, until recently, available membranes are invariably isotropic or homogeneous in structure. Such membranes are highly susceptible to pore plugging, even at the large cross flow velocities employed. Proper equipment configuration is a very essential factor for successful cell harvesting and it appears that this factor may have been overlooked in some designs. Thus pleated membrane cartridges or cassettes in which the membranes are held apart by a meshed screen are prone to channel blockage caused by the accumulation of cells or protein aggregates inside the mesh. It should be noted that during the process of concentration, the viscosity of the suspension may change by a factor of 500-1,000, which represents an overwhelming challenge to the design engineer. In this paper the authors report the use of a new cross flow microfiltration system which is capable of handling cell suspensions of over 20% dry weight. A cost comparison between the membrane and centrifugal processes is also discussed. Experimental
Cross flow unit The cross flow microfiltration unit, in the form of a prototype, held 1 m* of a membrane with an asymmetric structure. The membrane itself was supplied under the tradename Asypor TM. The production units can hold up
309
to 5 m2 of membrane; these are expected to be commercially available by September 1984. Both the membrane and the cross flow units were produced by Domnick Hunter Filters Ltd. Cell culture
Erwinia carotovora (NCPPB 1066) is grown in 500-l batches [l]. For the purpose of the experiment 100 1 of broth at a concentration of 0.55% dry weight was withdrawn from the fermenter and processed immediately. Results and discussion Cell harvesting Figure 1 shows the flux profiles for 3 membranes of different nominal pore sizes over a period of 4 hours. Flux declines sharply at the beginning, then more gradually after the first hour. The decrease in flux with time is due both to fouling and to the effect of the increase in the bulk concentration. It is interesting to note that the 0.6 pm membrane exerts some 20% more flux than the 0.2 pm membrane, but the 0.45 pm membrane shows almost twice the amount of flux exerted by the 0.6 pm membrane. Such an effect has been ascribed to the variation between the relative size of the bacterial cell to the pore and the distance between the pores [ 71; for instance, if the particle diameter is greater than the distance between two pores then simultaneous blocking of two adjacent pores is impossible because of steric hindrance. This shows that optimal pore size cannot be predicted from first principles for a particular microorganism at this stage but rather some experimentation is required. 60 p Erwn,a
carotovora
Im2crossflow Asypor
membrane
? =280 kPa
cell har,es’ins
Prototype
cnit
ITSUI
T= 3’“C
Time (hour; Fig. 1. The rate of flux decline
during
cell harvesting.
As the retentate becomes depleted of water, the cell concentration increases (as illustrated by Fig. 2). The rate of concentration at the start was extremely slow, because the quantity of solid was constant, so that at dilute concentrations a large amount of water had to be removed to make a significant change in the retentate concentration. However, at the high concentration end a small change in the water volume caused a large change in the cell concentration. Problems are likely to be encountered as the cell concentration approaches the 20-25% dry weight limit. In this concentration range the cell suspension was a thick slurry with an extremely high viscosity, which presented considerable difficulty to pumping due to the high pressure developed in the unit and the pipeline. Positive displacement pumps, particularly the low shear Archimedian screw type, are best suited-for this application. 25 Lellrecovery wlfh 20 -
0
Initial P:200
the crosstlow
Prototype
vol.=1001 @ 0.55 % dry kPa T= 37 C”
Feed velocity
= 69cm/s
2
1 Time
unit
wt
3
4
5
(hour)
Fig. 2. The rate of concentration
in cell harvesting.
The flux is plotted against the retentate concentration in Fig. 3, showing a semi-log relationship. There are two distinct regions in which the mass transfer coefficients appear constant. At very low cell concentrations the mass transfer coefficient is very high (25 l/m’-hr approx.). The transition occurs as the concentration reaches about 1% dry weight where the mass transfer coefficient is reduced by a factor of 5. The reason for this is unclear. Fane et al. [8] found that similar plots for activated sludge were curves concaving upwards. These workers pointed out that since they operated at a constant stirring rate, the mass transfer coefficient should decrease as the concentration increased, which they observed for their plots. However, if the mass transfer coefficient decreases with increasing concentration, one would expect the flux values to lie below a line of constant slope and, therefore, the actual plot should concave downward.
311
Cell
Recovery
Asypor P=BO
with
the
membranes kPa
T=
crassi&
prototype
unit
(TSU)
37’
C
L=69
cm/s
i;‘-1f,, ‘..
--
0 -1
I
I
I
I
0
1
2
3
--
‘1
4 ---
‘,
h
I
5
In (Cell cone) Fig. 3. Flux
concentration
dependence
for different
membrane
pore sizes.
The effect of pump shear on cells The effect of shear stress on bacterial cells has been observed, although what happens to the bacteria under this influence is still a matter of conjecture. It has been reported that under high cross flow velocity some strains of E. coli lost their ability to agglutinate yeast cells and this was suggested to be owing to the shearing off the bacteria surface of some specific proteins [3]. Some bacteria also lose their ability to adhere to glass or metal surfaces if they are subjected to recirculation in a centrifugal pump although their viability seems unaffected [ 91. Other bacteria lose their colloidal properties and settle under gravity (M.S. Le, unpublished results). In cell harvesting it is essential that the cell wall remains intact. The effect of pump shear on cells has been investigated as follows. A suspension of Erwinia cells at 1% dry weight was recirculated in a gear pump at a rate of 10 passes per minute, and at regular time intervals samples of the suspension were taken. Alkali was added to the samples to lyse the cells and the viscosity of the sample measured. DNA molecules are highly susceptible to degradation by shear stresses and they are released when cells are broken. Since these materials have a very high viscosity effect, measuring the lysate viscosity gave a good indication of the amount of whole cells remaining at any particular time. It has been found that under these experimental conditions fresh cells were not disrupted by pump shear. However, cells which have been frozen and thawed were considerably weakened. Table 1 shows the lysate viscosity after the cells had been recirculated in the pump for periods. It is clear that the rate of cell damage is a function of time, which presumably means it is a function of the concentration of intact cells. Figure 4 shows a typical profile of the
312 TABLE 1 Lysate viscosity after the whole cells (Erwinia carotovora, recirculated in a gear pump at 10 passes per minute Recirculation (min)
Viscosity (mPa-see)
time
at 300 sec.’
at 1200 set-’
0
9.68
5 15 45 60 75
5.18 4.48 4.48 2.36 1.89
8.02 6.61 5.19 6.0 4.72 4.25
37 ti
1% dry weight) have been
Erwinia
carotovora
(cells frozen/thawed
1% dry wt in 0.25M potassium Alkalrne in a gear
200
lysis after
Ihour
phaspha+e
once) pH6.5
of reclr:ularnn
pump @ lOpasses/m~nute
400
Shear rate Cl/s) Fig. 4. Rheological
behaviour of cell lysate.
lysate viscosity against the viscometer shear rate. The profile shows both pseudoplastic behaviour at low shear rates and dilatant behaviour at the high shear rates. The pseudoplastic property can be attributed to proteins and nucleic materials, whereas dilatant behaviour is probably due to the cell walls (M.S. Le, unpublished work; also Refs. [ 121 and [ 131). Thus, as shown in Table 1, a large change in the viscosity at the lower shear rate is an indication of the degradation of the DNA and a smaller change in the viscosity at the high shear rate means that the cell walls are less destructable.
313
Cell debris removal Figures 5-7 show the flux, enzyme transmission and protein transmission, respectively, for cells that have been lysed by alkali. In general, both the flux and the transmission are considerably lower than for lysate produced by the 30%lysed p:220
ceils
I” EDTA pH6 3 bet wf om~s)
kPa
T=26”C
3 45,~ Asypor
membrane
ICSUI
x 3
iz
5 l
01 0
’
’
20
8
40 Time
Fig. 5. Flux variation
9cm/s
. I
60
100
80
GO
140
160
(m’ln)
in the microfiltration
100
of cell lysate at various feed velocities.
?O%lwetwt; lysed 0.45~ AsyporiOSL)
ceils I” ECITA pH6 P- 220 kPa T= 26°C
27 cm/s
Timelmini Fig. 6. Time
variation
180
of enzyme
transmission
in cell lysate microfiltration.
314 3o"b(wet wt! Iyse? ce!lsin 8 mM EDTA pH6 3.~5~ ~sypor !osu] P=220 hFa T=26"C
130 cm/s
70.'
'+ L+
+
iCicm/s +
.
0
20
40
60
80
160
180
Time (mini Fig. 7. Time variation with protein transmission.
lysozyme disruption method [lo]. This can be explained by the fact that the alkaline disruption method causes protein denaturation and aggregation. The pump causes the protein precipitates to break up, exposing highly hydrophobic groups which go to build up the fouling layer [ 111, which reduces both the flux and the enzyme transmission. It should be noted that, in this case, different membrane pore sizes have not been explored, The 0.45 pm Asypor membrane produces a very clear permeate (turbidity of 0.05, compare{ to the centrifugal supernatant (Eppendorf 5415 microcentrifuge, 6 minutes at 998Og), which has a turbidity of 0.20). Other pore sizes could yield a much greater flux and high transmission at a still acceptable turbidity level.
Process economics Using the information shown in Fig. 3 process scale-up was performed by way of a digital computer simulation. Figure 8 shows some typical concentration profiles generated by the computer for various membrane areas used in the harvesting of a 500-l batch of cell suspension. It shows that 7.5 m2 would be required to concentrate 500 1 of broth at 0.55% dry weight to 22% dry weight in a period of 2% hours. A centrifugal separator (model KA25, Westfalia Separator Ltd) operating at 6500 rpm currently used for routine production on our pilot plant at the Centre for Applied Microbiology & Research performs the same duty in about 3 hours. In the scale-up process using membrane if all else being constant, the membrane area required is directly proportional to the batch size. Thus in the above example if the batch size is increased to 1000 1 the total membrane required
315
1
htlal
;
15.00 :
vol=5001 @0.55%
V=69cm/s .LS_L Asypor AZ membrane
dry wt
membrane area W)
Time
(hour)
Fig. 8. The rate of concentration
in cell harvesting for different
membrane
area sizes.
would be 15 m2. In the case of plate and frame equipment extra plates may be added to increase the membrane area until the capacity limit of the frame is reached. Further increase in the number of plates requires additional frames. Scaling up with centrifuges is somewhat less flexible and this is performkd either by employing a larger model or small models in multiple. Table 2 shows all the assumptions for the cost comparison between the centrifugal process and the membrane process for cell harvesting. In a biotechnological process, the recovery time is important because of the sensitive nature of some products. Our practice currently allows a maximum of 3 hours for each fermentation batch (500 1) plus 2 hours for cleaning and sanitisation. The number of batches require processing averages about 200 per year. For the cost analysis exercise we consider three ranges of scale of the operation. It should be noted that in a process of this nature, the cell recovery equipment is idle for a major part of the time. A large batch (5000 1) could be subdivided into smaller batches and processed by smaller size equipment which could then be fully utilized. In practice, small fermenters are often used in multiple instead of using the larger fermenters. Fermentation is then staggered in such a way that downstream processing equipment is ready when a fermenter requires to be processed. We compare two cases where on the one hand a large batch requires processing in a short time and on the other the same batch is subdivided either by physically dividing the batch volume or by staggering the fermentation with a series of smaller fermenters. The labour requirement merits some elaboration. In our experience, running a small membrane unit requires less than 50% of an operator’s time. For a membrane plant with 10 small modules it, is estimated that one operator on full time should be sufficient to control the process. On the other hand for each operating centrifuge there must be one operator on duty.
TABLE
w
2
Assumptions
for a cost comparison
between
various processes
Parameters
Centrifugal
Batch size (1)
l-300
4-900
25,000
40,000
CSA7 a
CSAaa
Capital cost including auxiliary equipment Basic system Membrane Membrane
($)
or model
7.5
Energy cost (Z/kW-hr) Labour cost (di/hr) Annual maintenance (g) Operating time (hr/yr)
0.05 3
Overheads, Vupplied
processes
period
(yr)
% of total cost by Westfalia
Separator
harvesting Membrane
70
processes
5,000
l-300
4-900
150,000
40,000
7,000
5 x CSAga
CSA8 a
lX6m’
5,000
life (yr) cost (g/m’)
Total power rating (kW)
Depreciation
for Erwinia carolouora
50
0.05 3
IO
0.05 3
1,500
2,000
8,000
1,000
1,000
1,000
0.05 3 4,000 4,800
5,000
5.000
20,000
80,000
20,000
3x5mZ
10
2 300
2 300
1
3
0.05 3 300 1,000
0.05 3 500 1,000
X
10 mZ
2 300
3X5m2 1 300 3
10
0.05 3
0.05 3 1,500
1,000
1,000 4,800
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
Ltd (UK).
TABLE
3
Annual
operating
cost comparison
Cost item (6)
between
the centrifugal
Centrifugal batch size
processes,
l-300
4-900
Membrane replacement Energy Labour Depreciation Overheads Annual Maintenance
375 3,000 2,500 825 1,500
Total
8,200
500 3,000 4,000 1,100 2,000 10,600
and membrane
processes
for Erwinia carotouora
Membrane batch size
cell harvesting
processes,
5,000
5,000
l-300
4--900
5,000
2,500 15,000 15,000 4,500 8,000
2,400 14,400 4,000 2,800 4,000
900 50 1,500 700 400 300
2,250 150 1,500 2,000 700 500
15,000 500 3,000 8,000 3,100 1,500
45,000
27,600
3,850
7,100
31,100
5,000 4,500 720 7,200 2,000 1,780 1,000 17,200
318
Table 3 shows the annual operating costs for each process. It is clear that for a centrifugal process the labour cost is one of the largest cost items (3050%) and this increases linearly with the number of centrifuges in the process. It also increases proportionally to the number of operating hours. Energy requirement seems to be a function of the batch size only. On the other hand with a membrane process membrane replacement is the largest cost item, accounting for up to 50% of the running cost. Labour cost increases slowly with the operation scale, but it varies in proportion with the number of operating hours. Energy consumption accounts for less than 5% of the total running cost. Overall, membrane processes are 30-50% less expensive to run than centrifugal processes. It appears that the cost differential is greatest on the small scale (up to 300 1). Within the range 500-5000 1 the cost differential between the two processes seems to be unaffected by the batch size. It is very interesting that on a large scale the cost of cell recovery is dramatically reduced if the operation is performed in smaller batches by practising staggered fermentation. By breaking a 500-I batch down to five 1000-l batches the cost is reduced by 39% and 45% for a centrifugal and a membrane process, respectively. Conclusion A new design of cross flow microfiltration equipment using asymmetric microporous membranes capable of handling cell concentrations of over 20% dry weight has been found suitable for harvesting Erwinia carotouora and subsequent separation of the cell debris. Cell lysis by pump shear is not a problem with fresh cells, although cells which have been frozen and thawed are considerably more fragile and require gentler treatment. A cost comparison for membrane processes and centrifugal processes for cell harvesting shows that the former are 30-5070 less expensive to run than the latter. By staggering the cell recovery process cost savings of up to 45% could be achieved. Acknowledgements The authors are grateful to Professor A. Atkinson for the provision of facilities, support and critical discussion, and our thanks are due to Mr. C.T. Billiet for the enthusiasm he has shown in this research project and for his encouragement in the preparation of this paper. References 1
K.A. Cammack, D.I. Marlborough and D.S. Miller, Physical properties and subunits structure of L-asparaginase isolated from Erwinia carotouora, Biochem. J., 126 (1972) 361.
319 J.D. Henry and R.C. Allred, Concentration of bacterial cells by crossflow filtration, Dev. Ind. Microbial., 13 (1972) 177. G.B. Tanny, D. Mirelman and T. Pistole, Improved filtration technique for concentrating and harvesting bacteria, Appl. Environ. Microbial., 40 (1980) 269. C.S. Genovesi, Several uses for tangential flow filtration in the pharmaceutical industry, J. Parenteral Sci. Tech., 37 (1983) 81. D.E. Reid and C. Adlam, Large scale harvesting and concentration of bacteria by tangential flow filtration, J. Appl. Bacterial., 41 (1974) 321. G.B. Tanny, D. Hauk and U. Merin, Biotechnical applications of a pleated cross flow microfiltration module, Desalination, 41 (1982) 299. MS. Le and J.A. Howell, An alternative model for ultrafiltration, Paper presented at IMTEC ‘83 Symposium, Sydney, November 1983. A.R. Cooper (Ed.), Ultrafiltration membranes and applications, Plenum Press, New York, NY, 1980, pp. 631-658. H. Fowler, Dept. Chemical Engineering, University College of Swansea, personal 10
11
12 13
communication, 1982. M.S. Le, L.B. Spark and P.S. Ward, The separation of aryl acylamidase by cross flow microfiltration and the significance of enzyme/cell debris interaction, J. Membrane Sci., 21(1984) 219. M.S. Le and J.A. Howell, The fouling of UF membranes and its treatment, in: Progress in Food Engineering, Forster Verlag AG/Foster Publishing Ltd., Switzerland, 1983, pp. 321-326. A. Leduy, A study of the rheological properties of a non-Newtonian fermentation broth, Biotech. Bioeng., 16 (1974) 61. N. Blakebrough, Rheology measurements on Asp. niger fermentation systems, J. Appl Chem. Biotech., 28 (1978) 453.
of Membrane
Elsevier
Science
MICROBIAL
M.S.
21(1984)
B.V.,
Amsterdam
ASPARAGINASE
Hunter
Filters
Ltd.,
307 in The Netherlands
RECOVERY
Durham
BY MEMBRANE
PROCESSES*
Road, Birtley.
Co. Durham,
DH3 2SF (Great Britain)
SPARK
Microbial Technology Laboratory, Porton Down, Salisbury, Wiltshire, P.S.
307-319 - Printed
LE**
Domnick L.B.
Science,
Publishers
WARD
Domnick (Received
PHLS Centre for Applied SP4 OJG (Great Britain)
Microbiology
& Research,
and N. LADWA
Hunter
Filters Ltd.,
December
19,
1983;
Durham accepted
Road,
Birtley,
in revised
Co. Durham,
form
April
DH3 2SF (Great Britain)
12, 1984)
Summary A feasibility study in the use of membranes for the recovery of L-asparaginase from the fermentation stage has been undertaken. In the cell harvesting stage 100 1 of broth (0.55% dry weight approx) of Erwinia carotovora is concentrated to 3.2 1 by a 1 m’ prototype cross flow microfiltration unit in just over 4 hours. Cell disruption was carried out by raising the cell suspension to pH 11.5 to release the enzyme, which is stable in alkali; for a short period. After neutralisation by acetic acid addition, the enzyme is separated from the cell debris by filtration with 0.45 pm Asypor membranes. A preliminary economic analysis shows that the membrane process can compete favourably with the conventional process, which employs high speed centrifuges.
Introduction L-asparaginase is frequently used in the treatment of some forms of leukaemia [l]. This therapeutic enzyme is an intracellular product of the organism Erwinia carotouoru (NCPPB 1066). Currently the microorganism is grown in batches of 500 1 to a concentration of approximately 0.55% dry weight. Cell harvesting is carried out by a continuous flow centrifugal separator which concentrates the broth by about 36-fold. The cell paste obtained from this process is usually frozen and stored for a short period before enzyme extraction is performed. Since L-asparaginase is stable at highly alkali pH [ 11, cell disruption is conveniently carried out by the addition of alkali to the cell suspensions to pH 11.5 for about 20 minutes. The nucleic acid and other cell debris materials which are released from the disruption process form a thick, *Paper presented at IMTEC ‘83, November 8-10, 1983, **To whom all correspondence should be addressed,
0376-7388/84/$03.00
0 1984
Elsevier
Science
Sydney,
Publishers
Australia.
B.V.
308
jelly-like mass, which is removed by the addition of acetic acid, which reduces the lysate pH to 6.5. Cell debris which includes partially dissolved cell walls and aggregates of protein precipitates is removed again by centrifugation at high g forces (usually >lO,OOOg). The use of large centrifuges has many disadvantages, both from the economic and safety viewpoints. In the first instance, centrifugal installation entails very high capital costs. A good deal of supervision is required during running and the frequent bowl change is a necessary process which is tedious and labour intensive. From the safety angle, where cost cannot be measured in terms of money and where opinion can vary considerably, the operation of centrifuges can be dangerous, with the possible risk of disintegration. Specially enclosed housings and procedures must be designed to deal with pathogenic organisms. This is extremely difficult in large-scale processing. Cross flow microfiltration, particularly with the new generation of high performance asymmetric microporous membranes, should minimise the operating costs, eliminate the danger of disintegration and go a long way towards reducing the contamination hazard. Cross flow microfiltration has been used for bacterial cell harvesting for over a decade [2], mainly on a laboratory scale where only small batches of up to 10 1 of culture suspension have been concentrated to about 200 ml [3,4]. Several claims of large-scale bacterial harvesting by cross flow membrane systems have been made. In all cases, however, the final cell concentrations achieved were typically less than 2% dry weight [ 5,6]. There are several factors which limit the final concentration of the cells. Firstly, until recently, available membranes are invariably isotropic or homogeneous in structure. Such membranes are highly susceptible to pore plugging, even at the large cross flow velocities employed. Proper equipment configuration is a very essential factor for successful cell harvesting and it appears that this factor may have been overlooked in some designs. Thus pleated membrane cartridges or cassettes in which the membranes are held apart by a meshed screen are prone to channel blockage caused by the accumulation of cells or protein aggregates inside the mesh. It should be noted that during the process of concentration, the viscosity of the suspension may change by a factor of 500-1,000, which represents an overwhelming challenge to the design engineer. In this paper the authors report the use of a new cross flow microfiltration system which is capable of handling cell suspensions of over 20% dry weight. A cost comparison between the membrane and centrifugal processes is also discussed. Experimental
Cross flow unit The cross flow microfiltration unit, in the form of a prototype, held 1 m* of a membrane with an asymmetric structure. The membrane itself was supplied under the tradename Asypor TM. The production units can hold up
309
to 5 m2 of membrane; these are expected to be commercially available by September 1984. Both the membrane and the cross flow units were produced by Domnick Hunter Filters Ltd. Cell culture
Erwinia carotovora (NCPPB 1066) is grown in 500-l batches [l]. For the purpose of the experiment 100 1 of broth at a concentration of 0.55% dry weight was withdrawn from the fermenter and processed immediately. Results and discussion Cell harvesting Figure 1 shows the flux profiles for 3 membranes of different nominal pore sizes over a period of 4 hours. Flux declines sharply at the beginning, then more gradually after the first hour. The decrease in flux with time is due both to fouling and to the effect of the increase in the bulk concentration. It is interesting to note that the 0.6 pm membrane exerts some 20% more flux than the 0.2 pm membrane, but the 0.45 pm membrane shows almost twice the amount of flux exerted by the 0.6 pm membrane. Such an effect has been ascribed to the variation between the relative size of the bacterial cell to the pore and the distance between the pores [ 71; for instance, if the particle diameter is greater than the distance between two pores then simultaneous blocking of two adjacent pores is impossible because of steric hindrance. This shows that optimal pore size cannot be predicted from first principles for a particular microorganism at this stage but rather some experimentation is required. 60 p Erwn,a
carotovora
Im2crossflow Asypor
membrane
? =280 kPa
cell har,es’ins
Prototype
cnit
ITSUI
T= 3’“C
Time (hour; Fig. 1. The rate of flux decline
during
cell harvesting.
As the retentate becomes depleted of water, the cell concentration increases (as illustrated by Fig. 2). The rate of concentration at the start was extremely slow, because the quantity of solid was constant, so that at dilute concentrations a large amount of water had to be removed to make a significant change in the retentate concentration. However, at the high concentration end a small change in the water volume caused a large change in the cell concentration. Problems are likely to be encountered as the cell concentration approaches the 20-25% dry weight limit. In this concentration range the cell suspension was a thick slurry with an extremely high viscosity, which presented considerable difficulty to pumping due to the high pressure developed in the unit and the pipeline. Positive displacement pumps, particularly the low shear Archimedian screw type, are best suited-for this application. 25 Lellrecovery wlfh 20 -
0
Initial P:200
the crosstlow
Prototype
vol.=1001 @ 0.55 % dry kPa T= 37 C”
Feed velocity
= 69cm/s
2
1 Time
unit
wt
3
4
5
(hour)
Fig. 2. The rate of concentration
in cell harvesting.
The flux is plotted against the retentate concentration in Fig. 3, showing a semi-log relationship. There are two distinct regions in which the mass transfer coefficients appear constant. At very low cell concentrations the mass transfer coefficient is very high (25 l/m’-hr approx.). The transition occurs as the concentration reaches about 1% dry weight where the mass transfer coefficient is reduced by a factor of 5. The reason for this is unclear. Fane et al. [8] found that similar plots for activated sludge were curves concaving upwards. These workers pointed out that since they operated at a constant stirring rate, the mass transfer coefficient should decrease as the concentration increased, which they observed for their plots. However, if the mass transfer coefficient decreases with increasing concentration, one would expect the flux values to lie below a line of constant slope and, therefore, the actual plot should concave downward.
311
Cell
Recovery
Asypor P=BO
with
the
membranes kPa
T=
crassi&
prototype
unit
(TSU)
37’
C
L=69
cm/s
i;‘-1f,, ‘..
--
0 -1
I
I
I
I
0
1
2
3
--
‘1
4 ---
‘,
h
I
5
In (Cell cone) Fig. 3. Flux
concentration
dependence
for different
membrane
pore sizes.
The effect of pump shear on cells The effect of shear stress on bacterial cells has been observed, although what happens to the bacteria under this influence is still a matter of conjecture. It has been reported that under high cross flow velocity some strains of E. coli lost their ability to agglutinate yeast cells and this was suggested to be owing to the shearing off the bacteria surface of some specific proteins [3]. Some bacteria also lose their ability to adhere to glass or metal surfaces if they are subjected to recirculation in a centrifugal pump although their viability seems unaffected [ 91. Other bacteria lose their colloidal properties and settle under gravity (M.S. Le, unpublished results). In cell harvesting it is essential that the cell wall remains intact. The effect of pump shear on cells has been investigated as follows. A suspension of Erwinia cells at 1% dry weight was recirculated in a gear pump at a rate of 10 passes per minute, and at regular time intervals samples of the suspension were taken. Alkali was added to the samples to lyse the cells and the viscosity of the sample measured. DNA molecules are highly susceptible to degradation by shear stresses and they are released when cells are broken. Since these materials have a very high viscosity effect, measuring the lysate viscosity gave a good indication of the amount of whole cells remaining at any particular time. It has been found that under these experimental conditions fresh cells were not disrupted by pump shear. However, cells which have been frozen and thawed were considerably weakened. Table 1 shows the lysate viscosity after the cells had been recirculated in the pump for periods. It is clear that the rate of cell damage is a function of time, which presumably means it is a function of the concentration of intact cells. Figure 4 shows a typical profile of the
312 TABLE 1 Lysate viscosity after the whole cells (Erwinia carotovora, recirculated in a gear pump at 10 passes per minute Recirculation (min)
Viscosity (mPa-see)
time
at 300 sec.’
at 1200 set-’
0
9.68
5 15 45 60 75
5.18 4.48 4.48 2.36 1.89
8.02 6.61 5.19 6.0 4.72 4.25
37 ti
1% dry weight) have been
Erwinia
carotovora
(cells frozen/thawed
1% dry wt in 0.25M potassium Alkalrne in a gear
200
lysis after
Ihour
phaspha+e
once) pH6.5
of reclr:ularnn
pump @ lOpasses/m~nute
400
Shear rate Cl/s) Fig. 4. Rheological
behaviour of cell lysate.
lysate viscosity against the viscometer shear rate. The profile shows both pseudoplastic behaviour at low shear rates and dilatant behaviour at the high shear rates. The pseudoplastic property can be attributed to proteins and nucleic materials, whereas dilatant behaviour is probably due to the cell walls (M.S. Le, unpublished work; also Refs. [ 121 and [ 131). Thus, as shown in Table 1, a large change in the viscosity at the lower shear rate is an indication of the degradation of the DNA and a smaller change in the viscosity at the high shear rate means that the cell walls are less destructable.
313
Cell debris removal Figures 5-7 show the flux, enzyme transmission and protein transmission, respectively, for cells that have been lysed by alkali. In general, both the flux and the transmission are considerably lower than for lysate produced by the 30%lysed p:220
ceils
I” EDTA pH6 3 bet wf om~s)
kPa
T=26”C
3 45,~ Asypor
membrane
ICSUI
x 3
iz
5 l
01 0
’
’
20
8
40 Time
Fig. 5. Flux variation
9cm/s
. I
60
100
80
GO
140
160
(m’ln)
in the microfiltration
100
of cell lysate at various feed velocities.
?O%lwetwt; lysed 0.45~ AsyporiOSL)
ceils I” ECITA pH6 P- 220 kPa T= 26°C
27 cm/s
Timelmini Fig. 6. Time
variation
180
of enzyme
transmission
in cell lysate microfiltration.
314 3o"b(wet wt! Iyse? ce!lsin 8 mM EDTA pH6 3.~5~ ~sypor !osu] P=220 hFa T=26"C
130 cm/s
70.'
'+ L+
+
iCicm/s +
.
0
20
40
60
80
160
180
Time (mini Fig. 7. Time variation with protein transmission.
lysozyme disruption method [lo]. This can be explained by the fact that the alkaline disruption method causes protein denaturation and aggregation. The pump causes the protein precipitates to break up, exposing highly hydrophobic groups which go to build up the fouling layer [ 111, which reduces both the flux and the enzyme transmission. It should be noted that, in this case, different membrane pore sizes have not been explored, The 0.45 pm Asypor membrane produces a very clear permeate (turbidity of 0.05, compare{ to the centrifugal supernatant (Eppendorf 5415 microcentrifuge, 6 minutes at 998Og), which has a turbidity of 0.20). Other pore sizes could yield a much greater flux and high transmission at a still acceptable turbidity level.
Process economics Using the information shown in Fig. 3 process scale-up was performed by way of a digital computer simulation. Figure 8 shows some typical concentration profiles generated by the computer for various membrane areas used in the harvesting of a 500-l batch of cell suspension. It shows that 7.5 m2 would be required to concentrate 500 1 of broth at 0.55% dry weight to 22% dry weight in a period of 2% hours. A centrifugal separator (model KA25, Westfalia Separator Ltd) operating at 6500 rpm currently used for routine production on our pilot plant at the Centre for Applied Microbiology & Research performs the same duty in about 3 hours. In the scale-up process using membrane if all else being constant, the membrane area required is directly proportional to the batch size. Thus in the above example if the batch size is increased to 1000 1 the total membrane required
315
1
htlal
;
15.00 :
vol=5001 @0.55%
V=69cm/s .LS_L Asypor AZ membrane
dry wt
membrane area W)
Time
(hour)
Fig. 8. The rate of concentration
in cell harvesting for different
membrane
area sizes.
would be 15 m2. In the case of plate and frame equipment extra plates may be added to increase the membrane area until the capacity limit of the frame is reached. Further increase in the number of plates requires additional frames. Scaling up with centrifuges is somewhat less flexible and this is performkd either by employing a larger model or small models in multiple. Table 2 shows all the assumptions for the cost comparison between the centrifugal process and the membrane process for cell harvesting. In a biotechnological process, the recovery time is important because of the sensitive nature of some products. Our practice currently allows a maximum of 3 hours for each fermentation batch (500 1) plus 2 hours for cleaning and sanitisation. The number of batches require processing averages about 200 per year. For the cost analysis exercise we consider three ranges of scale of the operation. It should be noted that in a process of this nature, the cell recovery equipment is idle for a major part of the time. A large batch (5000 1) could be subdivided into smaller batches and processed by smaller size equipment which could then be fully utilized. In practice, small fermenters are often used in multiple instead of using the larger fermenters. Fermentation is then staggered in such a way that downstream processing equipment is ready when a fermenter requires to be processed. We compare two cases where on the one hand a large batch requires processing in a short time and on the other the same batch is subdivided either by physically dividing the batch volume or by staggering the fermentation with a series of smaller fermenters. The labour requirement merits some elaboration. In our experience, running a small membrane unit requires less than 50% of an operator’s time. For a membrane plant with 10 small modules it, is estimated that one operator on full time should be sufficient to control the process. On the other hand for each operating centrifuge there must be one operator on duty.
TABLE
w
2
Assumptions
for a cost comparison
between
various processes
Parameters
Centrifugal
Batch size (1)
l-300
4-900
25,000
40,000
CSA7 a
CSAaa
Capital cost including auxiliary equipment Basic system Membrane Membrane
($)
or model
7.5
Energy cost (Z/kW-hr) Labour cost (di/hr) Annual maintenance (g) Operating time (hr/yr)
0.05 3
Overheads, Vupplied
processes
period
(yr)
% of total cost by Westfalia
Separator
harvesting Membrane
70
processes
5,000
l-300
4-900
150,000
40,000
7,000
5 x CSAga
CSA8 a
lX6m’
5,000
life (yr) cost (g/m’)
Total power rating (kW)
Depreciation
for Erwinia carolouora
50
0.05 3
IO
0.05 3
1,500
2,000
8,000
1,000
1,000
1,000
0.05 3 4,000 4,800
5,000
5.000
20,000
80,000
20,000
3x5mZ
10
2 300
2 300
1
3
0.05 3 300 1,000
0.05 3 500 1,000
X
10 mZ
2 300
3X5m2 1 300 3
10
0.05 3
0.05 3 1,500
1,000
1,000 4,800
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
Ltd (UK).
TABLE
3
Annual
operating
cost comparison
Cost item (6)
between
the centrifugal
Centrifugal batch size
processes,
l-300
4-900
Membrane replacement Energy Labour Depreciation Overheads Annual Maintenance
375 3,000 2,500 825 1,500
Total
8,200
500 3,000 4,000 1,100 2,000 10,600
and membrane
processes
for Erwinia carotouora
Membrane batch size
cell harvesting
processes,
5,000
5,000
l-300
4--900
5,000
2,500 15,000 15,000 4,500 8,000
2,400 14,400 4,000 2,800 4,000
900 50 1,500 700 400 300
2,250 150 1,500 2,000 700 500
15,000 500 3,000 8,000 3,100 1,500
45,000
27,600
3,850
7,100
31,100
5,000 4,500 720 7,200 2,000 1,780 1,000 17,200
318
Table 3 shows the annual operating costs for each process. It is clear that for a centrifugal process the labour cost is one of the largest cost items (3050%) and this increases linearly with the number of centrifuges in the process. It also increases proportionally to the number of operating hours. Energy requirement seems to be a function of the batch size only. On the other hand with a membrane process membrane replacement is the largest cost item, accounting for up to 50% of the running cost. Labour cost increases slowly with the operation scale, but it varies in proportion with the number of operating hours. Energy consumption accounts for less than 5% of the total running cost. Overall, membrane processes are 30-50% less expensive to run than centrifugal processes. It appears that the cost differential is greatest on the small scale (up to 300 1). Within the range 500-5000 1 the cost differential between the two processes seems to be unaffected by the batch size. It is very interesting that on a large scale the cost of cell recovery is dramatically reduced if the operation is performed in smaller batches by practising staggered fermentation. By breaking a 500-I batch down to five 1000-l batches the cost is reduced by 39% and 45% for a centrifugal and a membrane process, respectively. Conclusion A new design of cross flow microfiltration equipment using asymmetric microporous membranes capable of handling cell concentrations of over 20% dry weight has been found suitable for harvesting Erwinia carotouora and subsequent separation of the cell debris. Cell lysis by pump shear is not a problem with fresh cells, although cells which have been frozen and thawed are considerably more fragile and require gentler treatment. A cost comparison for membrane processes and centrifugal processes for cell harvesting shows that the former are 30-5070 less expensive to run than the latter. By staggering the cell recovery process cost savings of up to 45% could be achieved. Acknowledgements The authors are grateful to Professor A. Atkinson for the provision of facilities, support and critical discussion, and our thanks are due to Mr. C.T. Billiet for the enthusiasm he has shown in this research project and for his encouragement in the preparation of this paper. References 1
K.A. Cammack, D.I. Marlborough and D.S. Miller, Physical properties and subunits structure of L-asparaginase isolated from Erwinia carotouora, Biochem. J., 126 (1972) 361.
319 J.D. Henry and R.C. Allred, Concentration of bacterial cells by crossflow filtration, Dev. Ind. Microbial., 13 (1972) 177. G.B. Tanny, D. Mirelman and T. Pistole, Improved filtration technique for concentrating and harvesting bacteria, Appl. Environ. Microbial., 40 (1980) 269. C.S. Genovesi, Several uses for tangential flow filtration in the pharmaceutical industry, J. Parenteral Sci. Tech., 37 (1983) 81. D.E. Reid and C. Adlam, Large scale harvesting and concentration of bacteria by tangential flow filtration, J. Appl. Bacterial., 41 (1974) 321. G.B. Tanny, D. Hauk and U. Merin, Biotechnical applications of a pleated cross flow microfiltration module, Desalination, 41 (1982) 299. MS. Le and J.A. Howell, An alternative model for ultrafiltration, Paper presented at IMTEC ‘83 Symposium, Sydney, November 1983. A.R. Cooper (Ed.), Ultrafiltration membranes and applications, Plenum Press, New York, NY, 1980, pp. 631-658. H. Fowler, Dept. Chemical Engineering, University College of Swansea, personal 10
11
12 13
communication, 1982. M.S. Le, L.B. Spark and P.S. Ward, The separation of aryl acylamidase by cross flow microfiltration and the significance of enzyme/cell debris interaction, J. Membrane Sci., 21(1984) 219. M.S. Le and J.A. Howell, The fouling of UF membranes and its treatment, in: Progress in Food Engineering, Forster Verlag AG/Foster Publishing Ltd., Switzerland, 1983, pp. 321-326. A. Leduy, A study of the rheological properties of a non-Newtonian fermentation broth, Biotech. Bioeng., 16 (1974) 61. N. Blakebrough, Rheology measurements on Asp. niger fermentation systems, J. Appl Chem. Biotech., 28 (1978) 453.